Gasoline fraction



Dec. 7, 1954 c. HEMMINGER GASOLINE FRACTION 2 Sheets-Sheet 1 Filed Dec.

@Fzar'les Hemmzlrzser Savanna:-

dbuoraag Dec. 7, 1954 c. HEMMINGER ,6

GASOLINE FRACTION Filed D86. 21, 1951 2 Sheets-Sheet 2 Cfzcrls Hemm zinger Savexzbov QEDL Clbbovrzag United States Patent GASOLINE FRACTION Charles E. Hemminger, Westfield, N. J., assignor to Standand Oil Development Company, a corporation of Dela- Ware Application December 21, 1951, Serial No. 262,763

4- Claims. (CL- 196-49).

The present invention is-fully disclosed in: the following specification and claims, reference. being had tothe accompanying drawing.

More particularly, the present invention relates to improvements in manufacturing gasoline from heavy naphtha stocks and is. directed toward producing a gasoline of improved volatility and road octane number.

v Briefly, the invention involves a combination process in which heavy naphtha, is. split into two streams, one of which is subjected to hydroforming-while the heavier portion of the naphtha is first subjected. to thermal cracking and thereafter treated under conditions conducive to hydrogenation of olefins in the cracked product, before the remaining portion of the feed is. subjected to hydroforming. In the preferred: modification of the invention the saturation of olefins in. the cracked fraction and the hydroforming. process are performed inthe same zone, and in the presence. of the same catalyst. This hydroforming. and saturation of the olefins from the cracking operation are performed in the presence of a fluidized bed of catalyst, under conditions so controlled that the olefins are not subjected to temperaturessutficiently high to cause degradation. This is particularly important when treating. a feedstock containing significant amounts of diolefin or other highly reactive and polymerizable olefinic constituents. While the tendency of such materials to polymerize and dehydrogenate further to. coke on the active catalyst surface can be completely inhibited by operating at extremely high hydrogen pressures, the intermediate and low pressure operations characteristic of hydroforming do not entirely prevent such coke formation, and this tendency is lessened by avoiding high operating temperatures. Gasoline volatility is an important fuel property, controlling. vapor lock tendencies, starting andwarm-up performance, as well asgasoline mileage. and other characteristics of the motor fuel. In meeting the requirements of refining for increased volatility, other workers in hydroforming have chosen special types of catalysts, or have operated at a relatively high hydroforming pressure in order to give hydrccracking of heavy naphtha feeds.

However, when hydrocracking. is attained in this way, in

order to produce the desirable increase in the relative amounts of the product boiling below 212 F, excessive amounts of light hydrocarbons such as propane and butane are also ordinarily formed. These have less than feed value, and are therefore undesirable products. Moreover, even the C5 and Cs hydrocarbons which are produced in this method of obtaining gasoline volatility are essentially saturated hydrocarbons and contain 40-60% straight chain normal paraffins These norrjnal paraffins have very low octane number, and the fraction containing them has only on the order of 75 or at the most 85 Research octane number. Thus, the final product of hydroforming with excessive hydrocracking to obtain volatility, results in a gasoline with relatively low octane number light hydrocarbons. Even though the heavy hydrocarbons have a compensating high octane number, this is an undesirablecombination for an automotive fuel. A motor fuel whose constituents contain hydrocarbons having the same octane number throughout the boiling range is much more satis factory; inother words, the light ends, the middle fractions and the heavy ends should all possess substantially the same laboratory octane ratings for optimum road cctane performance.

One object of the present invention is to convert a heavy naphtha under conditions which will produce an 2: improved gasoline product characterized in having a high roadv octane quality.

Another object of the present invention is to provide a process for converting naphtha into high quality gasoline as regards octane quality and volatility which process is so operated as to avoid excessive formation of light hydrocarbons.

Another object of the present, invention. is to provide a process adapted to convert virgin. naphtha into a motor fuel of improved octane quality and volatility without causing the formation of excessive quantities of carbonaceous degradation products.

Another object of the present invention is to improve the quality of refinery naphtha, as for example, that produced from a viscosity reducing operation, or a coking operation, to at least partly saturate the olefins therein contained to permit further conversion, without causing excessive formation of carbonaceous degradation products, followed by subsequent hydroforming toimprove the octane quality of such refinery naphtha.

A more specific object of the present invention is to hydroform naphtha. in a hydroforming zone under conditions such that a substantial portion of the heat necessary to support this reaction is supplied by accompanying exothermic reactions, as for example, the hydrogenation of olefins, both operations. being performed concurrently in the presence of added hydrogen and in. contact with the same catalyst.

Another object of the present invention is to saturate olefinic naphthas under closely controlled conditions of temperature so as to prevent cracking or other degradation of said olenfinc naphtha.

In carrying the foregoing objects and similar ends into practical effect, the present invention contemplates in one embodiment, a system in which heavy virgin naphtha is fractionated into a light fraction which is subjected to hydroforming, and a heavier fraction which is thermally cracked or reformed and thereafter subjected. to closely controlled conditions of hydrogenation, wherein the originm light fraction and the hydrogenated cracked fraction are combined and hydroformed. to form a motor fuel of improved road octane quality and volatility.

In the accompanying drawings, there is shown in Fig. I an apparatus layout in which a preferred modification of the present invention may be carried into effect; and in Fig. II there is shown a modification in, which the treating vessel wherein the cracked naphtha is subjected to a hydrogenation treatment to at least partially saturate the olefins therein contained forms also a part of the hydroforming reaction zone.

Similar reference characters refer to similar parts throughout the views.

Referring in detail to Fig. I, a virgin naphtha having a nominal boiling range of, say, 200-450 F. is introduced to the system through line 1,. thence charged to a fractional distillation column 2. From fractionator 2, a fraction boiling within the range of from about 175 to 320 or 350 F. is taken off overhead via line 3, forced through a heating coil 5 where it is heated to a temperature of approximately 1000-1100 F. and thereafter charged via line 6 to a reactor 7.

Reactor 7 contains a fluidized bed of catalyst C. This catalyst may be, for example, active alumina impregnated with or carrying a VI group metal oxide. For example, the catalyst may contain of active alumina and 10% by weight of molybdenum oxide. The catalyst carrier may be a spinel such as zinc aluminate, or alumina containing a promotional amount of an activator such as zinc oxide, say, 1040 Weight per cent of the total weight of the alumina plus molybdena. The catalyst is composed preferably of particles of mixed sizes within the range of, say, 20 to 80 microns. The reactor 7 is provided with the usual grid or screen G through which the incoming hydrogen-containing gas is forced in order to provide good distribution.

Referring again to fractional distillation column 2, the bottoms, that is, the material boiling above from 320 to 350? F. up to about 450 F. is withdrawn from the said fractionating column 2, through line 8, forced through a pump P, and thence into a thermal reforming or cracking unit 9, which is operated under a pressure of about 750 p. s. i. g. The naphtha vapor is maintained in unit 9 under a temperature of about 1000l100 F. for a sufficient period of time to convert it to the extent of about 10-17 Weight per cent to C3 and lighter hydrocarbons.

The cracked product stream is withdrawn from cracking unit 9 through line 10, and forced through a cooling means 11 where it is cooled to a temperature sutficient to condense material boiling above about 400 F. The cooled stream is then passed via line 12 into a separator 13 from the bottom of which the condensed material is withdrawn through line 14. The uncondensed vapor fraction is Withdrawn overhead from separator 13 through line 15, and charged to a fractional distillation column 16. Normally gaseous hydrocarbons withdrawn overhead from column 16 through line 17 are conducted to a conventional absorber 18, where known methods are used to separate C3 and C4 hydrocarbons from the lighter hydrocarbons in the product stream. The C2 and lighter product gases are withdrawn overhead from absorber 18 via line 19. C3 and C4 hydrocarbons are withdrawn from the bottom via line 20 and after separation from absorber medium in a still not shown are charged to a catalytic polymerization unit 21 where they are treated under known conditions to form polymer gasoline. This gasoline is recovered via line 22, and delivered to storage with other products of the present system as hereinafter more fully explained.

Referring again to fractional distillation column 16, an intermediate fraction boiling within the range of from about C to 200 F. is withdrawn as a side stream through line 23 and mixed with the polymer gasoline in line 22 and delivered with the latter to product storage.

The bottoms fraction from fractional distillation column 16 is withdrawn through line 24. this heavy cracked naphtha material boiling substantially within the range of from about ZOO-400 F., as previously indicated. This material may also be admixed with an extraneous supply of olefinic naphtha supplied through line 25, such as a cracked naphtha derived from the viscosity reducing or coking of a heavy hydrocarbon oil. In either event, the mixed or unmixed cracked naphtha material in line 24 is charged to a furnace or other heating means 26, wherein it is ra idly heated to a temperature of from about 600-900 F. The reheated cracked naphtha or olefinic naphtha is withdrawn from heating means 26 throu h line 27 and charged to reactor 28. Meanwhile,

a h drogen-containing gas recovered from the product purification system, which may be relatively cool, is fed from line 29 into line 30 and thereafter into the bottom of reactor 28. Reactor 28 has disposed therein a grid or screen G, through which the incoming hydrogencontaining gas is forced. Catalyst withdrawn at reaction temperature from reactor 7 throu h an aerated standpipe 31 is also charged into reactor 28 at a point above the grid G1 but in close proximity thereto, the flow of catalvst through standpipe 31 being controlled by valve 32. The gas flow passes upwardly with the catalyst suspended therein through reactor 28 at a superficial velocity of about /2-1 ft. per second. In reactor 28 the olefins contained in the hydrocarbons are at least partially saturated, particularly the most active and least stable constituents thereof. Thereafter the hydrogenated naphtha, the hydrogen-containing gas and the catalyst are With drawn overhead from 28 via line 33 and charged into reactor 7, entering the bed of catalyst C at an upper level. above the inlet of line 6 feeding the hot naphtha fraction to the reactor.

The total gasiform material entering vessel 7 is also caused to flow upwardly therethrough at /21 ft. per second superficial velocity, so as to maintain the bed of catalyst C in a well fluidized state. The bed of catalyst, as usual, extends from grid G to an upper dense phase level L. Above L in reactor 7 there is a relatively light dispersed phase in which the concentration of catalyst in gasitorm material decreases upwardly by the settling out of suspended particles. Within the upper portion of reactor 7 there is also disposed a plurality of solidsseparating devices 23 (one shown) through which the gasiform material about to exit from the reactor is forced, for the purpose of separating entrained catalyst. Separated catalyst is returned to the dense phase by dip pipes d (one shown).

The product withdrawn overhead from reactor 7 via. line 34 is passed through a cooling means 35 and discharged at operating pressure into a separator 36. The normally liquid constituents of the product condensed in separator 36 are Withdrawn via line 37 and charged to a stabilizer 38. From stabilizer 38 mixed C-irhYCllO- carbons are withdrawn overhead through line 39. The bottoms from stabilizer 38 are withdrawn through line 40 and charged to a rerun still 41, from which the finished C5-430 F. hydroformed gasoline is withdrawn overhead through line 42. This product is mixed with the polymer gasoline and light cracked naphtha in line 22 and delivered to storage 43. The bottoms from rerun still 41 are withdrawn through line 44 and may be mixed with heating oil or otherwise utilized as chemicals or as chemical raw materials.

Referring again to separator 36, the uncondensed material which is primarily hydrogen is withdrawn overhead via line 45. A portion of this gas may be rejected from the system as a purge stream through line 46, together with impurities such as gaseous sulfur compounds, light hydrocarbons, nitrogen, etc. The main stream is recycled to the system via line 47 and a compressor 48 from which the compressed hydrogen-containing gas is conducted via line 49 to reactors 7 and 28, respectively. As previously indicated, with respect to the recycled gas in line 49, a portion of this may be charged virtually unheated via line 29 to reactor 28. On the other hand, the recycled gas passing to reactor 7 is forced through heater 50 before it is fed to the bottom of reactor 7. The temperature of the hydrogen-containing gas fed to reactor 28 may be controlled, if desired, by varying the proportions of cool gas entering line 30 from line 29 and hot gas applied from line 51 via line 52, depending on the tem perature conditions to be maintained in reactor 28. Primarily, the gas fed to reactor 28 is a cold recycled gas, and by the same token, the recycled hydrogen gas fed to reactor 7 must be hot gas. The reason for these gas temperatures is that the reaction in 28 is primarily exothermic, Whereas that in reactor 7 is primarily endothermic.

The catalyst in reactor 7 with the passage of time will require regenerati n due to the deposition thereon of carbonaceous material. In regenerating the catalyst in reactor 7, it is withdrawn through an aerated standpipe 54, thence through line 55 to regeneration vessel 56 located at an upper level. The rate of catalyst withdrawal through standpipe 54 is controlled by a slide valve 57, and a stripping gas such as steam is introduced via line 58 just above valve 57 which serves the double function of aerating the catalyst in standpipe 54 and stripping entrained hydrocarbons from this stream of catalyst prior to the regeneration process. A portion of the regeneration air supply, or other oxygen-containing gas, introduced through line 59 picks up the catalyst entering line 55 from standpipe 54 and carries it up into vessel 56. Additional regeneration air, which may be the main supply of oxygen-containing gas to the regeneration process, is introduced into vessel 56 via line 60.

The total gas supply to vessel 56 is controlled at such a rate, preferably within the range of from about /2 to 1 ft. per second, that the catalyst within this vessel is maintained in the form of a dense turbulent suspension. Vessel 56 is arranged in such a position that the level L of the turbulent bed of catalyst within vessel 56 is at all times well above the level L of catalyst bed C in reactor 7. By virtue of this arrangement, regenerated catalyst flowing from vessel 56 is returned by gravity to a lower level of catalyst bed C.

A cooling coil 61 may also be supplied in the bed of catalyst in vessel 56, if desired, to avoid developing excessive catalyst temperatures during the regeneration process. The desired temperature level may be within the range of from about 1100 to 1300 F., preferably about 1200 F. Regeneration fumes are removed from vessel 56 via line 62, which may include a cyclone separator or other auxiliary equipment (not shown) to return entrained catalyst particles from these vent gases into the circulating catalyst stream.

Regenerated catalyst is Withdrawn downwardly from vessel 56 via aerated standpipe or withdrawal line 63. The catalyst particles in this stream are subjected to a short-time, high-temperature stripping process in vessel 65, which may be of any suitable design. Stripping gas for this process is a hot hydrogen-containing stream introduced via line 66, which branches ofi from the main line 51 feeding hydrogen recycle gas to reactor 7. This Feed Stocln Hydroformiug Stage:

its-m lieu hot: hydrogen-containing: gas= removes residual negen elation gaszfrom: the catalyst stream; and may also give: it at controlledl but; short hydrogen: pretreatment before entering the main reactor: The: solids. residence time; in stripper 65: may be veryshort, oh the order-"of 110 seconds on lhssa so -that essentially only stripping and: little such treating takes: place. Spent gas fronn thisstripping process-passesdromi vessel 65' through line G'Z 'di'rectIy' i'ntioi the reactor so as: to simplify the questiom of? catalyst re covery atthis -points. ILino 67 enters vessel: 7:' the dispurse phase: portion above bed level: Stripped catalyst firomwessel 65 fiowsdown. by gravity" through line: 68; entering the: dense'fiuid bed: ofi catalyse C. in reactor. 7

a: point just above griiii G- While the-duawing hasashowm spent catalyst withdrawn only from. reactor 7 and catalyst in treating. vessel: 28 going only overhead into: the: reactor 7 it may; be ad vantageous in. certain cases: to: withdraw spent: catalyst from vessel 281 through: a bottom. drawoflr line not shown andcirculate this catalyst through; line 55 to; regenerate! establishing; a; separate pretrcating Z0132. Ina the: lower portion. of; bed. 6., whether; onnot a: hafile: B.- is; employed this hot hydrogen-containing gas and. hot naphtha intro: duced through. line 6,. undergo. theendothermic. hydro.- formingreactiom under the influence of hot: regenerated catalyst which has, been withdrawn through line 54. and returned: through line; 68 as." in: Figure: 1. The upward flowing: stream: of; catalyst and reactant gases, cooled this reaction, then comes in contact with the olefinie naphtha: stream; introduced through line 33%. At this point, the exothermic hydrogenation of the. olefinic con; stituentsin this crackedinaphtha. stream takcsplace under the influence of the hydroforming: catalyst and hydrogen in the; gas stream; In: this: manner, the temperature of the; reactant mixture which may be. as. much as 2:5 lower at the point where the olefinic; feed enters through line 33' is again raised so that the average temperature is maintained at approximately the same level in both the upper and lower portions of catalyst bed C.

Having thus described the operation of the process in vessel 56 Sucha circulation might: be in. addition. to or general, the invention. may be. better understood i t rms entirel y' replacing the: withdrawal of spent: catalyst" from of the; following specific data. as to. operating conditions: Table I- i Example I Example II 2'VesselSystem, Figure 1 One Split Vessel; Figure 2 Operating Range:v

Preferred Cnnditiious v Operatingv Range. Preferred Conditions 1 Heavy Virgin Naphtha 'Boiling Range E 1757450 7 0. N. AS'IM Vapor Pressure, R. V; P-.. Hydrogenationstager Feed Stock;. F

Qat./oi1'.weight.ratio.-... Hydrogen-rate, CFIE M01. percent H2. Inlet; Temperature, F. Pressure, 1;. s. 1. Outlet Temperature, Product: percent ,unsats Feed rate, w./lir Catalyst 10% MolybdenaiomAluemina Oxide, Gel Cat. loll weight ratio. 33 n- Hydrogen rate, (BF/B.

Moll. percent Hr Temperature,, E Pressure, 9. s. i: g

Product- 0''. N. (Res. Clear);

Heavy Virgin Naphthay. 175-450 From hydrofonming zone. in vessel. 7..

Heavy Virgin; ap h 175-425.

- Heavy Virgin N aphtha.

350 to FBP.

170-20. From hydroior-rnin-g zo ne in vessel 7 ..s-3.0-. 'Frorm hydroforming I zone Luvessel 7..

.ol'fi "2'.

IB'PtO 350-400 IB'Pto-35 0. 0.23.0; 0134.0. 10%.M'01ybdena on Alu- 10% Molybdena on Aluna Ox de 91-.

mina Oxide Gel spec-4,000.

1 Poundsot oil per hour per pound of catalyst in thoreaction-spaeo. Cubic test; per barrel! of'oil; 1 Based. on total feed.

reactor 7.. The choice here will be determined by the character of the. feed stock. and the treating. reaction in vessel 28-, andthe tendency ofi carbonaceous deposits. to

be laid down at this point as compared to. the hydroforming; reactor 7.

Refernowto Figure. 2,. there is illustrated. here a fed throughline 2.4 to preheater 26 contains relatively little of the more highly unstabl yp t. fin y rocarbons. Sucha naphtha stream, which may include catalytically cracked naphtha or other cracked naphthas containing moderate amounts of unsaturated and some naphthen-ic constituents, can be fed into the main hydroforrning reactor without the preliminary hydrogentation step shown in vessel 28 in Figure 1. In this case: the inlet line 33 feeding olefinic naphtha to reactor 7 leads directly from the preheating furnace 25, Reactor 7" may, if desired, be modified slightly by including. a horizontal or perforated bafile plate; id at an intermediate level within the bed of catalyst. C. The function of such a bafiieis to" decrease the relative amount of back-mixing ofi catalyst, and reactant gas: from that portion of the fluid :bed which lies above the bathe to that portion which lies tom of reactor 7' via line 51'" without the necessity for.-

It will be seen that in the process as carried out in either Figure 1' or Figure 2 the exothermic heat of the hydrogenation of ol'efinic constituents in the cracked naphtha stream is employed to offset the endothermic heat of" reaction of the hydroforming process. The separate hydrogen pretreating vessel 28, in Figure 1, permits the use of more highly unsaturated and unstable cracked naphthas in this system than otherwise could be employed without causing high amounts of coke to be deposited on the catalyst; The olefi'uic feeds for which this. type of pretreatment isparti'cularly helpful, will ordinarily be those which are low innaphthene content, which means that they would contain a relatively low amount of' available hydrogen to be released in the hydroforming process. The amount of prehydrogenation which the cracked naphtha in line 24- und'ergoes in this embodiment of the present invention depends upon the relative amounts of hydrogen cracked naphtha fed to vessel 28, and is variable depending upon the composition of the naphtha; teed.

On the other hand, the apparatus of Figure 2, as indicated above, is ordinarily more useful when employing cracked naphtha feeds containing less olefin, or more particularly less of the unstable olefi-ns, and containing ordinarily somewhat more of the naphthenic, constituents. The deciding factor as to. whether or not aseparate hydroen pretr at nt y he pl yed to. advantage is thus de rmine to some exten by he. relativ amounts o olefi an naphth u hydrocarb n i he r k ap h str am.

from 5055% 212 In the simplified design of Figure 2, the arrangement providing the introduction of the olefin feed in the upper portion of the catalyst bed permits the use of relatively high inlet temperatures in the range of about 1000-1200 F. for the catalyst, hydrogen and naphtha streams entering the bottom of the reactor 7'. At the same time, the cracked naphtha stream entering at the upper level can be fed at a somewhat lower temperature, of from about 600900 F. This minimizes the tendency for feed degradation, while still utilizing the principle of direct mixing of the hydrogenation and hydroforming reactions to even out the temperature drop ordinarily encountered in the hydroforming process.

The advantage of the method of producing a superior motor gasoline by the invention herein described is shown also in the following data:

The fuel No. 1 is produced according to the present invention whereas fuel No. 2 is produced according to previous practice. In the latter case the light virgin naphtha is blended directly with the total hydroformed heavy naphtha, giving a low octane number PEP-175 F. fraction. However, in fuel No. 1 the light naphtha is processed as described above, and its octane number improved from 86.0 to 89.5 Research 0. N. The two fuels were evaluated in the conventional Uniontown road octane rating method, evaluating road performance against standard reference fuels in a number of automobiles representing an average cross section of car makes. For a 1.4 increase in clear Research 0. N., a 3.0 point higher road rating is observed for fuel No. 1, made in accordance with the specification of the present invention. Expressed in another way, fuel No. 1 gives a road 0. N. only 3.8 points lower than the laboratory rating, while fuel No. 2 shows 5.4 points lower road octane number.

A still further advantage of the present invention lies in the fact that this desirable octane distribution is ac companied by desirable volatility characteristics in the motor fuel produced. As already indicated above, ordinary methods of producing a more volati e gasoline, such as blending in various refinery C4 or C4-C5 streams or cracking for front end volatility, frequently give an unbalanced fuel composition with relatively too small an amount of material in the intermediate light naphtha boiling range. Stated more specifically, when such a fuel is blended to lbs. Reid vapor pressure wi h butane. it ordinarily contains more material b ilin ofi be w l53 F. than is desirable f r a given am unt b iling be w 212 F., by standard A. S. T. M. meth ds of test. The process of the resent invention makes possible the pre aration of a. hi h oualitv m 11-. D, V, P. motor f l having not more than about 15-20% ((0 158 F. with F. by Engler distillation, with the hydroformed light naphtha, light cracked fraction of the original naphtha and polymer gasoline all available to contribute high octane components in the intermediate boiling range.

It will be understood that the invention 1s not to be limited by the particular feed stocks or apparatus arrangements shown here by way of illustration, and that various modifications and applications of the same basic principle are within the purview of this invention limited only by the following claims.

What is claimed is:

1. The method of producing from heavy virgin naphtha a motor gasoline having improved road octane rating and volatility characteristics which comprises distilling said naphtha to produce a lighter fraction boiling within the range of from about 175 to 350 F. and a heavier fraction boiling within the range of from about 350-450 F., vaporizing and heating said lighter fraction to a temperature above about 1000" F. and charging said hot vapors to a reaction zone containing a fluidized bed of hydroforming catalyst and a recycled stream of hydrogencontaining gas under hydroforming conditions of temperature, pressure and contact time, subjecting said heavier fraction to thermal cracking conditions of temperature, pressure and contact time to form a cracked naphtha product stream including at least 10 weight per cent of Ca-hydrocarbons, separating from said cracked naphtha stream a C3-C4. olefinic product and converting the olefins therein to a polymer gasoline of intermediate boiling range and high octane rating, further separating from said cracked naphtha a light C5-200" F. naphtha of high octane rating and a heavy olefinic naphtha of approximately 200400 F. boiling range, rapidly heating a cracked naphtha stream including said last-named heavy olefinic fraction to a temperature of from about 600-900 F. and subjecting it to a mild partial hydrogenation employing hydrogen derived from the abovementioned hydroforming reaction and a bed of fluidized catalyst drawn directly from said hydroforming catalyst bed at a temperature below the inlet temperature of said lighter naphtha and catalyst to the hydroforming reaction zone, adding said partly hydrogenated stock to the fluidized reaction mixture of hydroforming catalyst and lighter naphtha and completing the simultaneous hydroforming of said lighter naphtha and hydrogenation of said olefinic naphtha in a common reaction zone comprising at least a part of said hydroforming zone, and combining said polymer gasoline, said light cracked naphtha and the product of said hydroforming and hydrogenation reactions to produce a naphtha having high octane quality in all boiling ranges and a volatility of less than 20% at 158 F. with at least 50% at 212 F. when blended with refinery butane to a 10 pound vapor pressure, by standard A. S. T. M. methods of test.

2. The method of producing from heavy virgin naphtha a motor gasoline having improved road octane rating and volatility characteristics, which comprises distilling said naphtha to produce a lighter fraction boiling within the range of from about -350 F. and a heavier fraction boiling within the range of about 350450 F., maintaining a hydroforming reaction zone with a fluidized stream of a hydroforming catalyst passing therethrough, introducing a vaporized stream of said lighter naphtha fraction into a dense bed of said catalyst maintained in said reaction zone under hydroforming conditions of temperature, pressure and hydrogen-containing gas atmosphere, recovering from said hydroforming zone a product stream including entrained catalyst, at least partly hydrof rmed naphtha of improved octane characteristics and additional hydrogen produced during the hydroforming reaction, subjecting said heavier naphtha fraction to thermal cracking conditions of temperature, pressure and contact time to form a cracked naphtha product stream including at least 10 weight per cent of (Es-hydrocarbons, recovering from said cracked naphtha a heavy olefinic fraction of approximately 200400 F. boiling range, rapidly heating a cracked naphtha stream including said last-named fraction to a temperature of from about 600-900 F. and subjecting it to a mild hydrogenation reaction employing a bed of fluidized catalyst drawn directly from the above-mentioned hydroforming reaction together with hydrogen derived from said hydroforming reacti n, combining the said partially hydroformed light naphtha and partially hydrogenated cracked heavy naphtha together with said catalyst and hydrogen in a common zone, and completing said hydroforming reaction and said hydrogenation reaction in said common zone wherein the exothermic hydrogenation reaction supplies at least a portion of the heat required for the endothermic hydroforming reaction.

3. The process of claim 2 in which the total reaction mixture of feed stock, catalyst, and hydrogen-containing gas in said hydroforming zone is conducted into the hydrogenation zone together with the olefinic feed stock undergoing hydrogenation therein.

4. The process according to claim 2 in which the total reaction mixture in said hydrogenation reaction, including catalyst, feed stock, and hydrogen-containing gas, is conducted into the hydroforming zone together with the feed stock undergoing hydroforming.

(References on following page) References Cited in the file of this patent Iglgl2l26i'65 UNITED STATES PATENTS 2:400:795

Number Name Date 2,426,870 2,154,064 Cooke Apr. 11, 1939 5 2,436,170 2,193,798 Atwell Mar. 19, 1940 2,644,785

2,289,716 Marschnerf July 14, 1942 10 Name Date Layng et a1. July 13, 1943 Watson May 21, 1946 Hill Sept. 2, 1947 Hill Feb. 17, 1948 Harding et a1. July 7, 1953 

1. THE METHOD OF PRODUCING FROM HEAVY VIRGIN NAPHTHA A MOTOR GASOLINE HAVING IMPROVED ROAD OCTANE RATING AND VOLATILITY CHARACTERISTICS WHICH COMPRISES DISTILLING SAID NAPHTHA TO PRODUCE A LIGHTER FRACTION BOILING WITHIN THE RANGE OF FROM ABOUT 175* TO 350* F. AND A HEAVIER FRACTION BOILING WITHIN THE RANGE OF FROM ABOUT 350*-450* F., VAPORIZING AND HEATING SAID LIGHTER FRACTION TO A TEMPERATURE ABOVE ABOUT 1000* F. AND CHARGING SAID HOT VAPORS TO A REACTION ZONE CONTAINING A FLUIDIZED BED OF HYDROFORMING CATALYST AND A RECYCLED STREAM OF HYDROGENCONTAINING GAS UNDER HYDROFORMING CONDITIONS OF TEMPERATURE, PRESSURE AND CONTACT TIME, SUBJECTING SAID HEAVIER FRACTION TO THERMAL CRACKING CONDITIONS OF TEMPERATURE, PRESSURE AND CONTACT TIME TO FORM A CRACKED NAPHTHA PRODUCT STREAM INCLUDING AT LEAST 10 WEIGHT PER CENT OF C3-HYDROCARBONS, SEPARATING FROM SAID CRACKED NAPHTHA STREAM A C3-C4 OLEFINIC PRODUCT AND CONVERTING THE OLEFINS THEREIN TO A POLYMER GASOLINE OF INTERMEDIATE BOILING RANGE AND HIGH OCTANE RATING, FURTHER SEPARATING FROM SAID CRACKED NAPHTHA A LIGHT C5-200* F. NAPHTHA OF HIGH OCTANE RATING FROM A HEAVY OLEFINIC NAPHTHA OF APPROXIMAELY 200*-400* F. BOILING RANGE, RAPIDLY HEATING A CRACKED NAPHTHA STREAM INCLUDING SAID LAST-NAMED HEAVY OLEFINIC FRACTION TO A TEMPERATURE OF FROM ABOUT 600*-900* F. AND SUBJECTING IT TO A MILD PARTIAL HYDROGENATION EMPLOYING HYDROGEN DERIVED FROM THE ABOVEMENTIONED HYDROFORMING REACTION AND A BED OF FLUIDIZED CATALYST DRAWN DIRECTLY FROM SAID HYDROFORMING A CATALYST BED AT A TEMPERATURE BELOW THE INLET TEMPERATUR OF SAID LIGHTER NAPHTHA AND CATALYST TO THE HYDROFORMING REACTION ZONE, ADDING SAID PARTLY HYDROGENATED STOCK TO THE FLUIDIZED REACTION MIXTURE OF HYDROFORMING CATALYST AND LIGHTER NAPHTHA AND COMPLETING THE SIMULTANEOUS HYDROFORMING OF SAID LIGHTER NAPHTHA AND HYDROGENATION OF SAID OLEFINIC NAPHTHA IN A COMMON REACTION ZONE COMPRISING AT LEAST A PART OF SAID HYDROFORMING ZONE, AND COMBINING SAID POLYMER GASOLINE, SAID LIGHT CRACKED NAPHTHA AND THE PRODUCT OF SAID HYDROFORMING AND HYDROGENATION REACTIONS TO PRODUCE A NAPHTHA HAVING HIGH OCTANE QUALITY IN ALL BOILING RANGES AND A VOLATILITY OF LESS THAN 20% AT 158* F. WITH AT LEAST 50% AT 212* F. WHEN BLENDED WITH REFINERY BUTANE TO A 10 POUND VAPOR PRESSURE, BY STANDARD A. S. T. M. METHODS OF TEST. 